Process for concentration of antibodies and therapeutic products thereof

ABSTRACT

The present disclosure provides a process for concentrating proteins including an ultrafiltering, a diafiltering, and a second ultrafiltering sequence, at elevated temperatures, such as above about 30° C. The disclosure also includes a process for preparing highly concentrated antibody compositions, and highly concentrated antibody products.

RELATED APPLICATIONS

This application claims the priority benefit of U.S. ProvisionalApplication Ser. No. 60/609,092 filed Sep. 9, 2004, which isincorporated herein by reference in its entirety.

BACKGROUND

Methods for isolating, purifying, and concentrating biological materialsare known and include, for example, chromatography, ultrafiltration, andlyophilization, see generally, R. Hatti-Kaul et al., “DownstreamProcessing in Biotechnology,” in Basic Biotechnology, Chap. 9, pages187-211, 2nd ed., Cambridge University Press (2001). Processes formaking concentrated monoclonal antibody preparations for administrationto humans are known, see for example, U.S. Pat. No. 6,252,055, whichuses ultrafiltration and which re-circulates the resulting filtrate.

Some challenges associated with available antibody concentration methodsinclude, for example, low fluxes, long process times, large membraneareas, mechanical recovery yield and losses, operator-intensiveintervention or handling, low mass transfer rates, energyinefficiencies, and hydraulic pressure limits on concentrationequipment. These and other challenges can contribute to a high totalcost of manufacture and ultimately higher costs to therapeutic drugconsumers.

There is a need for improved processes for preparing highly concentratedprotein formulations, such as liquid antibody preparations andtherapeutic products thereof.

SUMMARY

In general terms, the present disclosure generally relates to processesfor concentrating proteins, such as processes for concentrating anantibody preparation, pharmaceutical formulations containing such apreparation, and there use in human therapy or animal therapy.

In embodiments the present disclosure provides processes for preparinghighly concentrated proteins, such as antibody preparations; andtherapeutic products prepared by the process, such as therapeuticantibody products. Accordingly, the present disclosure provides, aprocess for concentrating proteins comprising: a first ultrafiltering ofa first antibody preparation to provide a second antibody preparation; adiafiltering the second antibody preparation to provide a diafilteredintermediate antibody preparation; and a second ultrafiltering of thediafiltered intermediate antibody preparation to provide a thirdantibody preparation, wherein one or more of the first ultrafiltering,the second ultrafiltering, and the diafiltering are accomplished atelevated temperatures, for example, from about 30° C. to about 50° C.

The present disclosure also provides, in embodiments, a process forconcentrating proteins comprising: a first ultrafiltering of a firstprotein mixture to provide a second protein mixture; a diafiltering thesecond protein mixture to provide a diafiltered protein mixture; and asecond ultrafiltering of the diafiltered protein mixture to provide athird protein mixture, wherein one or more of the first ultrafiltering,the diafiltering, and the second ultrafiltering are accomplished at, forexample, about 45° C.

The present disclosure also provides, in embodiments, a highlyconcentrated antibody composition prepared by the above processes.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 illustrates an apparatus for accomplishing the preparativeprocess, in embodiments of the present disclosure.

FIGS. 2 through 17 illustrate various observed or measured processvalues over various phases or mode of the process, in embodiments of thepresent disclosure.

FIGS. 18 and 19 illustrate the effect of elevated temperature on productquality, in embodiments of the present disclosure.

FIGS. 20 and 21 illustrate the effect of elevated temperature onbioburden control, in embodiments of the present disclosure.

FIG. 22 illustrates the effect of elevated temperature on process fluxand process time, in embodiments of the present disclosure.

FIGS. 23 through 25 illustrates various observed or measured processvalues over various phases or mode of the scaled-up process, inembodiments of the present disclosure.

DETAILED DESCRIPTION

Various embodiments of the present disclosure will be described indetail with reference to drawings, if any. Reference to variousembodiments does not limit the scope of the invention, which is limitedonly by the scope of the claims attached hereto. Additionally, anyexamples set forth in this specification are not intended to be limitingand merely set forth some of the many possible embodiments for theclaimed invention.

The following are used, unless otherwise described:

“Ultrafiltering,” “ultrafiltration,” “ultrafiltered,” “UF,” and liketerms refer to, for example, using synthetic semi-permeable membranes,with appropriate physical and chemical properties, to discriminatebetween molecules in the mixture, primarily on the basis of molecularsize and shape, and accomplish separation of different molecules oraccomplish concentration of like molecules.

“Diafiltering,” “diafiltration,” “diafiltered,” “diafiltrating,” “DF,”and like terms refer to, for example, using an ultrafiltration membraneto remove, replace, or lower the concentration of salts or solvents fromsolutions or mixtures containing proteins, peptides, nucleic acids, orother biomolecules.

“Transmembrane pressure” or “TMP” refers to the average applied pressurefrom the feed to the filtrate side of the membrane calculated as TMP[bar]=[(P_(F)+P_(R))/2]−P_(f). where P_(F) is the feed pressure, P_(R)is the retentate pressure, and P_(f) is the filtrate pressure.

“Tangential flow filtration,” “cross flow filtration,” “TFF,” and liketerms refer to a mode of filtration in which the solute-containingsolution passes tangentially across the UF membrane and lower molecularweigh salts or solutes are passed through by applying pressure.

“Antibody” is used in the broadest sense and specifically covers intactmonoclonal antibodies, polyclonal antibodies, multispecific antibodies(e.g., bispecific antibodies) formed from at least two intactantibodies, and antibody fragments, so long as they exhibit the desiredbiological activity. An antibody is a protein generated by the immunesystem that is capable of recognizing and binding to a specific antigen.Described in terms of its structure, an antibody is a Y-shaped proteinconsisting of four amino acid chains, two heavy and two light. In asimplified model sufficient for this appeal, each antibody has primarilytwo regions: a variable region and a constant region. The variableregion, located on the ends of the arms of the Y, binds to and interactswith the target antigen. This variable region includes a complementarydetermining region (CDR) that recognizes and binds to a specific bindingsite on a particular antigen. The constant region, located on the tailof the Y, is recognized by and interacts with the immune system(Janeway, C., Travers, P., Walport, M., Shlomchik (2001) Immuno Biology,5th Ed., Garland Publishing, New York). A target antigen generally hasnumerous binding sites, also called epitopes, recognized by CDRs onmultiple antibodies. Each antibody that specifically binds to adifferent epitope has a different structure. Thus, one antigen may havemore than one corresponding antibody.

The basic 4-chain antibody unit is a heterotetrameric glycoproteincomposed of two identical light (L) chains and two identical heavy (H)chains (an IgM antibody consists of 5 of the basic heterotetramer unitalong with an additional polypeptide called J chain, and thereforecontain 10 antigen binding sites, while secreted IgA antibodies canpolymerize to form polyvalent assemblages comprising 2-5 of the basic4-chain units along with J chain). In the case of IgGs, the 4-chain unitis generally about 150,000 daltons. Each L chain is linked to an H chainby one covalent disulfide bond, while the two H chains are linked toeach other by one or more disulfide bonds depending on the H chainisotype. Each H and L chain also has regularly spaced intrachaindisulfide bridges. Each H chain has at the N-terminus, a variable domain(V_(H)) followed by three constant domains (C_(H)) for each of the α andγ chains and four C_(H) domains for μ and ε isotypes. Each L chain hasat the N-terminus, a variable domain (V_(L)) followed by a constantdomain (C_(L)) at its other end. The V_(L) is aligned with the V_(H) andthe C_(L) is aligned with the first constant domain of the heavy chain(C_(H)1). Particular amino acid residues are believed to form aninterface between the light chain and heavy chain variable domains. Thepairing of a V_(H) and V_(L) together forms a single antigen-bindingsite. For the structure and properties of the different classes ofantibodies, see e.g., Basic and Clinical Immunology, 8th edition, D.Stites, A. Terr and T. Parslow (eds.), Appleton & Lange, Norwalk, Conn.,1994, page 71 and Chapter 6.

The L chain from any vertebrate species can be assigned to one of twoclearly distinct types, called kappa and lambda, based on the amino acidsequences of their constant domains. Depending on the amino acidsequence of the constant domain of their heavy chains (C_(H)),immunoglobulins can be assigned to different classes or isotypes. Thereare five classes of immunoglobulins: IgA, IgD, IgE, IgG, and IgM, havingheavy chains designated α, δ, ε, ε, and μ, respectively. The γ and αclasses are further divided into subclasses on the basis of relativelyminor differences in C_(H) sequence and function, e.g., humans expressthe following subclasses: IgG1, IgG2, IgG3, IgG4, IgA1, and IgA2.

The term “variable” refers to the fact that certain segments of thevariable domains differ extensively in sequence among antibodies. The Vdomain mediates antigen binding and define specificity of a particularantibody for its particular antigen. However, the variability is notevenly distributed across the approximately 110-amino acid span of thevariable domains. Instead, the V regions consist of relatively invariantstretches called framework regions (FRs) of 15-30 amino acids separatedby shorter regions of extreme variability called “hypervariable regions”that are each 9-12 amino acids long. The variable domains of nativeheavy and light chains each comprise four FRs, largely adopting aβ-sheet configuration, connected by three hypervariable regions, whichform loops connecting, and in some cases forming part of, the β-sheetstructure. The hypervariable regions in each chain are held together inclose proximity by the FRs and, with the hypervariable regions from theother chain, contribute to the formation of the antigen-binding site ofantibodies (see Kabat et al., in Sequences of Proteins of ImmunologicalInterest, 5th Ed. Public Health Service, National Institutes of Health,Bethesda, Md. (1991)). The constant domains are not involved directly inbinding an antibody to an antigen, but exhibit various effectorfunctions, such as participation of the antibody in antibody dependentcellular cytotoxicity (ADCC).

The term “hypervariable region” when used herein refers to the aminoacid residues of an antibody which are responsible for antigen-binding.The hypervariable region generally comprises amino acid residues from a“complementarity determining region” or “CDR” (e.g., around about Kabatresidues 24-34 (L1), 50-56 (L2) and 89-97 (L3) in the V_(L), and aroundabout Kabat residues 31-35B (H1), 50-65 (H2) and 95-102 (H3) in theV_(H) (see Kabat et al., supra) and/or those residues from a“hypervariable loop” (e.g., around about Chothia residues 26-32 (L1),50-52 (L2) and 91-96 (L3) in the V_(L), and 26-32 (H1), 52A-55 (H2) and96-101 (H3) in the V_(H) (Chothia and Lesk, J. Mol. Biol, 196:901-917(1987)).

The term “monoclonal antibody” as used herein refers to an antibody froma population of substantially homogeneous antibodies, i.e., theindividual antibodies comprising the population are identical and/orbind the same epitope(s), except for possible variants that may ariseduring production of the monoclonal antibody, such variants generallybeing present in minor amounts. Such monoclonal antibody typicallyincludes an antibody comprising a polypeptide sequence that binds atarget, wherein the target-binding polypeptide sequence was obtained bya process that includes the selection of a single target bindingpolypeptide sequence from a plurality of polypeptide sequences. Forexample, the selection process can be the selection of a unique clonefrom a plurality of clones, such as a pool of hybridoma clones, phageclones or recombinant DNA clones. It should be understood that theselected target binding sequence can be further altered, for example, toimprove affinity for the target, to humanize the target bindingsequence, to improve its production in cell culture, to reduce itsimmunogenicity in vivo, to create a multispecific antibody, etc., andthat an antibody comprising the altered target binding sequence is alsoa monoclonal antibody of this invention. In contrast to polyclonalantibody preparations which typically include different antibodiesdirected against different determinants (epitopes), each monoclonalantibody of a monoclonal antibody preparation is directed against asingle determinant on an antigen. In addition to their specificity, themonoclonal antibody preparations are advantageous in that they aretypically uncontaminated by other immunoglobulins. The modifier“monoclonal” indicates the character of the antibody as being obtainedfrom a substantially homogeneous population of antibodies, and is not tobe construed as requiring production of the antibody by any particularmethod. For example, the monoclonal antibodies to be used in accordancewith the present invention may be made by a variety of techniques,including, for example, the hybridoma method (e.g., Kohler et al.,Nature, 256:495 (1975); Harlow et al., Antibodies: A Laboratory Manual,(Cold Spring Harbor Laboratory Press, 2nd ed. 1988); Hammerling et al.,in: Monoclonal Antibodies and T-Cell Hybridomas, 563-681, (Elsevier,N.Y., 1981)), recombinant DNA methods (see, e.g., U.S. Pat. No.4,816,567), phage display technologies (see, e.g., Clackson et al.,Nature, 352:624-628 (1991); Marks et al., J. Mol. Biol., 222:581-597(1991); Sidhu et al., J. Mol. Biol. 338(2):299-310 (2004); Lee et al.,J. Mol. Biol. 340(5):1073-1093 (2004); Fellouse, Proc. Nat. Acad. Sci.USA 101(34):12467-12472 (2004); and Lee et al., J. Immunol. Methods284(1-2):119-132 (2004), and technologies for producing human orhuman-like antibodies in animals that have parts or all of the humanimmunoglobulin loci or genes encoding human immunoglobulin sequences(see, e.g., WO 1998/24893; WO 1996/34096; WO 1996/33735; WO 1991/10741;Jakobovits, et al., Proc. Natl. Acad. Sci. USA, 90:2551 (1993);Jakobovits, et al., Nature, 362:255-258 (1993); Bruggemann, et al., Yearin Immuno., 7:33 (1993); U.S. Pat. Nos. 5,545,806; 5,569,825; 5,591,669(all to GenPharm); 5,545,807; WO 1997/17852; U.S. Pat. Nos. 5,545,807;5,545,806; 5,569,825; 5,625,126; 5,633,425; and 5,661,016; Marks, etal., Bio/Technology, 10: 779-783 (1992); Lonberg, et al., Nature 368:856-859 (1994); Morrison, Nature, 368: 812-813 (1994); Fishwild, et al.,Nature Biotechnology, 14: 845-851 (1996); Neuberger, NatureBiotechnology, 14: 826 (1996); and Lonberg and Huszar, Intern. Rev.Immunol., 13: 65-93 (1995).

“Chimeric” antibodies (immunoglobulins) have a portion of the heavyand/or light chain identical with or homologous to correspondingsequences in antibodies derived from a particular species or belongingto a particular antibody class or subclass, while the remainder of thechain(s) is identical with or homologous to corresponding sequences inantibodies derived from another species or belonging to another antibodyclass or subclass, as well as fragments of such antibodies, so long asthey exhibit the desired biological activity (U.S. Pat. No. 4,816,567;and Morrison, et al., Proc. Natl. Acad. Sci. USA 81:6851-6855 (1984)).Humanized antibody as used herein is a subset of chimeric antibodies.

“Humanized” forms of non-human (e.g., murine) antibodies are chimericantibodies which contain minimal sequence derived from non-humanimmunoglobulin. For the most part, humanized antibodies are humanimmunoglobulins (recipient or acceptor antibody) in which hypervariableregion residues of the recipient are replaced by hypervariable regionresidues from a non-human species (donor antibody) such as mouse, rat,rabbit or nonhuman primate having the desired specificity, affinity, andcapacity. In some instances, Fv framework region (FR) residues of thehuman immunoglobulin are replaced by corresponding non-human residues.Furthermore, humanized antibodies may comprise residues which are notfound in the recipient antibody or in the donor antibody. Thesemodifications are made to further refine antibody performance such asbinding affinity. Generally, the humanized antibody will comprisesubstantially all of at least one, and typically two, variable domains,in which all or substantially all of the hypervariable loops correspondto those of a non-human immunoglobulin and all or substantially all ofthe FR regions are those of a human immunoglobulin sequence although theFR regions may include one or more amino acid substitutions that improvebinding affinity. The number of these amino acid substitutions in the FRare typically no more than 6 in the H chain, and in the L chain, no morethan 3. The humanized antibody optionally also will comprise at least aportion of an immunoglobulin constant region (Fc), typically that of ahuman immunoglobulin. For further details, see Jones, et al., Nature321:522-525 (1986); Reichmann, et al., Nature 332:323-329 (1988); andPresta, Curr. Op. Struct. Biol. 2:593-596 (1992).

“Antibody fragments” comprise a portion of an intact antibody,preferably the antigen binding or variable region of the intactantibody. Examples of antibody fragments include Fab, Fab′, F(ab′)₂, andFv fragments; diabodies; linear antibodies (see U.S. Pat. No. 5,641,870,Example 2; Zapata, et al., Protein Eng., 8(10): 1057-1062 (1995));single-chain antibody molecules; and multispecific antibodies formedfrom antibody fragments.

Papain digestion of antibodies produces two identical antigen-bindingfragments, called “Fab” fragments, and a residual “Fc” fragment, adesignation reflecting the ability to crystallize readily. The Fabfragment consists of an entire L chain along with the variable regiondomain of the H chain (VH), and the first constant domain of one heavychain (C_(H)1). Each Fab fragment is monovalent with respect to antigenbinding, i.e., it has a single antigen-binding site. Pepsin treatment ofan antibody yields a single large F(ab′)₂ fragment which roughlycorresponds to two disulfide linked Fab fragments having divalentantigen-binding activity and is still capable of cross-linking antigen.Fab′ fragments differ from Fab fragments by having additional fewresidues at the carboxy terminus of the C_(H)1 domain including one ormore cysteines from the antibody hinge region. Fab′-SH is thedesignation herein for Fab′ in which the cysteine residue(s) of theconstant domains bear a free thiol group. F(ab′)₂ antibody fragmentsoriginally were produced as pairs of Fab′ fragments which have hingecysteines between them. Other chemical couplings of antibody fragmentsare also known.

The Fc fragment comprises the carboxy-terminal portions of both H chainsheld together by disulfides. The effector functions of antibodies aredetermined by sequences in the Fc region, which region is also the partrecognized by Fc receptors (FcR) found on certain types of cells.

“Fv” is the minimum antibody fragment which contains a completeantigen-recognition and -binding site. This fragment consists of a dimerof one heavy- and one light-chain variable region domain in tight,non-covalent association. From the folding of these two domains emanatesix hypervariable loops (3 loops each from the H and L chain) thatcontribute the amino acid residues for antigen binding and conferantigen binding specificity to the antibody. However, even a singlevariable domain (or half of an Fv comprising only three CDRs specificfor an antigen) has the ability to recognize and bind antigen, althoughat a lower affinity than the entire binding site.

“Single-chain Fv” also abbreviated as “sFv” or “scFv” are antibodyfragments that comprise the V_(H) and V_(L) antibody domains connectedinto a single polypeptide chain. Preferably, the sFv polypeptide furthercomprises a polypeptide linker between the V_(H) and V_(L) domains whichenables the sFv to form the desired structure for antigen binding. For areview of sFv, see Pluckthun in The Pharmacology of MonoclonalAntibodies, vol. 113, Rosenburg and Moore eds., Springer-Verlag, NewYork, pp. 269-315 (1994).

“About” modifying, for example, the quantity of an ingredient in thecompositions, concentration of an active, buffer volumes, diavolumes,pore size, apparent molecular, molecular weight cut-off, processtemperature, process time, yields, flow rates, pressures, bio-burdens,and like values, and ranges thereof, employed in the methods of theinvention, refers to variation in the numerical quantity that can occur,for example, through typical measuring and handling procedures used formaking concentrates or use solutions; through inadvertent error in theseprocedures; through differences in the manufacture, source, or purity ofthe ingredients employed to make the compositions or carry out themethods; and like considerations. The term “about” also encompassesamounts that differ due to aging of a composition with a particularinitial concentration or mixture. The term “about” also encompassesamounts that differ due to mixing or processing a composition with aparticular initial concentration or mixture. Whether or not modified bythe term “about” the claims include equivalents to the quantities.

“Consisting essentially of” refers to a process of obtaining aconcentrated protein composition or antibody composition that includesthe steps and ingredients listed in the claim, plus other steps andingredients that do not materially affect the basic and novel propertiesof the composition, such as a multiplicity of steps or buffer media.Ingredients that materially affect the basic properties of thecomposition and method of the present disclosure impart undesirablecharacteristics including, for example, bio-burden, such as theundesirable toxicity or irritability associated with contaminants.

The indefinite article “a” or “an” and its corresponding definitearticle “the” as used herein is understood to mean at least one, or oneor more, unless specified otherwise.

The present disclosure provides, in embodiments, the abovementionedprocesses and the concentrated antibody products thereof.

In embodiments of the present disclosure, the preparative processes andproducts thereof can be used in preparing highly concentrated antibodypreparations and similar preparations, such as purifying andconcentrating proteins or like substances from natural or syntheticsources, and which products can be useful for treating pathologicalconditions, such as asthma, cancer, psoriasis, inhibiting angiogenesis,and like pathological conditions.

In embodiments of the above-mentioned process for preparing highlyconcentrated antibody compositions of the disclosure, the followingfurther exemplifies how to make and use the preparative processes andproducts of the disclosure.

In embodiments of the present disclosure, there is provided a processfor preparing highly concentrated antibody compositions, for example,according to accomplishing the following steps in the order recited,comprising:

a first ultrafiltering of a first antibody preparation, having aconcentration of, for example, about 0.1 to about 10 grams per liter(g/L), to provide an second antibody preparation as the retentate,having a greater antibody concentration of, for example, about 10 toabout 50 grams per liter;

a diafiltering of the resulting second antibody preparation to provide adiafiltered intermediate antibody preparation as the retentate, havingabout the same concentration as the resulting second antibodypreparation retentate, that is, diafiltering to accomplish a bufferexchange at constant volume; and

a second ultrafiltering of the diafiltered intermediate antibodypreparation to provide a third antibody preparation as the retentate,having a greater antibody concentration of, for example, about 150 toabout 200 grams per liter.

The preparative processes of the disclosure can further comprise anoptional product recovery step or steps, for example, and as disclosedand illustrated herein.

In embodiments of the above-mentioned process of the disclosure, one ormore of the first ultrafiltering, the diafiltering, and the secondultrafiltering, can be accomplished at, for example, from about 30° C.to about 70° C. In embodiments, these steps can also be accomplished at,for example, from about 30° C. to about 50° C. In embodiments, thesesteps can also be accomplished at, for example, from about 35° C. toabout 50° C. In embodiments, these steps can also be accomplished at,for example, about 45° C., such as from about 45° C. plus or minus 5° C.Depending upon the type of antibody preparation, for processesaccomplished at temperatures above about 70° C., the preparation mayshow signs of deterioration, such as denaturation, agglomeration, andlike phenomena. For processes accomplished at temperatures below fromabout 30 to about 35° C., the flow rates are typically undesirably lowand process times are undesirably long, making the process at lowertemperatures less attractive for efficient commercial production.

In embodiments, the first antibody preparation can have an antibodyconcentration of, for example, from about 0.1 to about 100 grams perliter (g/L). The antibody concentration is, for example, a commonconcentration typically available from other preliminary protein orantibody purification steps or methods, such as, centrifugation,filtration, chromatography, and like procedures. The resulting secondantibody preparation obtainable from the first ultrafiltering can havean antibody concentration of, for example, from about 10 to about 50grams per liter, and for example, about 20 to about 40 grams per liter,such as 30 grams per liter. A range for the antibody concentration ofthe intermediate antibody preparation can depend upon, for example, abalance of factors, such as sample volume and sample flux achievablewith a particular buffer containing the second antibody preparation. Theintermediate antibody preparation can have an antibody concentration of,for example, about 25 to about 35 grams per liter and the third antibodypreparation can have an antibody concentration of, for example, fromabout 170 to about 200 grams per liter. The third antibody preparation,in embodiments, can have an antibody concentration of, for example, fromabout 50 to about 250 grams per liter, such as from of about 100 toabout 230 grams per liter, and from about 170 to about 200 grams perliter, such as 185 grams per liter.

It will be apparent to one skilled in the art, upon comprehending thepresent disclosure, that the intermediate antibody preparation and thirdantibody preparation comprise the same ultra-filtered retentate exceptfor, for example, differences in the antibody concentration resultingfrom the first and second ultrafiltering concentration, and differencesin the suspending buffer media resulting from the diafiltration bufferexchange. Thus, there is little, if any, compositional change, such asdegradation, of the target protein or antibody product, in embodimentsof the present disclosure.

Conventional ultrafiltration concentration methods can have generallygreater time and lesser through-put inefficiencies having considerablylonger process times such as several days to several weeks, processconsiderably smaller volumes, or both.

In embodiments, the protein concentration process of the disclosure canbe accomplished in, for example, from about 1 to 10 hours, preferably infrom about 2 to 5 hours, and more preferably in about 3 hours. Thepreferences favor higher flux through-put and smaller membrane areas.

In embodiments, the first ultrafiltering can be accomplished, forexample, in about 35 percent of the total process time. Thus, forexample, in a concentration and purification process of the disclosurewith about 3 hours total process time, the first ultrafiltering can beaccomplished in about 45 minutes. In embodiments, the secondultrafiltering can be accomplished, for example, in about 15 percent ofthe total process time. Thus, for example, in a process of thedisclosure with about 3 hours total process time, the secondultrafiltering can be accomplished in about 15 minutes. The diafilteringcan be accomplished, for example, in about 50 percent of the totalprocess time. Thus, for example, in a process of the disclosure withabout 3 hours total process time, the diafiltering can be accomplishedin from about 90 to about 120 minutes.

In embodiments, the first ultrafiltering and the second ultrafilteringcan be accomplished, for example, with an ultra-filter membrane having anominal pore size, or molecular weight cut-off, of about 5 to about 50kiloDaltons. Another suitable nominal pore size is, for example, fromabout 10 to about 40 kiloDaltons. Yet another suitable nominal poresize, or molecular weight cut-off, is about 30 kiloDaltons.

In embodiments, the first antibody preparation can contain, for example,an antibody having an apparent molecular weight of, for example, about100 to about 200 kiloDaltons. In other embodiments, the first antibodypreparation can contain an antibody having an apparent molecular weightof, for example, about 150 kiloDaltons, such as when the antibodypreparation comprises anti-IgE antibodies or IgE, see for example, U.S.Pat. No. 6,172,213 assigned to Genentech, Inc.

Other antibodies suitable for use in the present disclosure includecancer treating antibodies, see generally, for example: PCT/US02/19592;PCT/US01/20118; PCT/US01/25464; PCT/US01/26626; PCT/US02/28859;PCT/US02/41798; PCT/US02/12206; PCT/US03/11148; PCT/US02/12619; andPCT/US02/33050. Still other antibodies suitable for use in the presentdisclosure include an anti-CD20 antibody and like antibodies includinghuman, non-human, murine, hybrid, and chimeric forms. See for exampleU.S. Pat. No. 6,582,959 (VEGF) and U.S. Patent Application No.2002/0122797 A1 (human VEGF).

In embodiments, antibodies included within the scope of the disclosureinclude hybrid and recombinant antibodies (e.g., “humanized” and “human”antibodies) regardless of species of origin or immunoglobulin class orsubclass designation, as well as antibody fragments (for example, Fab,F(ab′)₂, and F_(v)). See U.S. Pat. No. 4,816,567; Mage and Lamoyi, inMonoclonal Antibody Production Techniques and Applications, 79-97,Marcel Dekker, Inc., New York, (1987).

Monoclonal antibodies may also be used and can be isolated from phageantibody libraries using the techniques described in Clackson et al(1991) Nature, 352:624-628 and Marks, et al. (1991) J. Mol. Biol.,222:581-597, for example. Monoclonal antibodies include “chimeric”antibodies in which a portion of the heavy and/or light chain isidentical with or homologous to corresponding sequences in antibodiesderived from a particular species or belonging to a particular antibodyclass or subclass, while the remainder of the chain(s) is identical withor homologous to corresponding sequences in antibodies derived fromanother species or belonging to another antibody class or subclass, aswell as fragments of such antibodies, so long as they exhibit thedesired biological activity (U.S. Pat. No. 4,816,567; and Morrison, etal. (1984) Proc. Natl. Acad. Sci. USA, 81:6851-6855). Chimericantibodies can include “primatized” antibodies comprising variabledomain antigen-binding sequences derived from a non-human primate (e.g.,Old World Monkey, Ape etc) and human constant region sequences.

Monoclonal antibodies are highly specific, being directed against asingle antigenic site. Furthermore, in contrast to polyclonal antibodypreparations that include different antibodies directed againstdifferent determinants (epitopes), each monoclonal antibody is directedagainst a single determinant on the antigen. In addition to theirspecificity, the monoclonal antibodies are advantageous in that they maybe synthesized uncontaminated by other antibodies. Thus, the modifier“monoclonal” indicates the character of the antibody as being obtainedfrom such a substantially homogeneous population of antibodies, i.e.,the individual antibodies comprising the population are identical exceptfor possible naturally occurring mutations that may be present in minoramounts, and is not to be construed as requiring production of theantibody by any particular method. For example, the monoclonalantibodies for use in the disclosure may be made using the hybridomamethod first described by Kohler & Milstein, Nature, 256:495 (1975), ormay be made by recombinant DNA methods. Other known methods of antibodyproduction are described, for example, in Goding, Monoclonal Antibodies:Principles and Practice, 59-103, Academic Press (1986); Kozbor, J.Immunol., 133:3001 (1984). Brodeur, et al., Monoclonal AntibodyProduction Techniques and Applications, 51-63, Marcel Dekker, Inc., NewYork (1987).

Various methods have been employed to produce monoclonal antibodies(MAbs). Hybridoma technology, which refers to a cloned cell line thatproduces a single type of antibody, uses the cells of various species,including mice (murine), hamsters, rats, and humans. Another method toprepare MAbs uses genetic engineering including recombinant DNAtechniques. Monoclonal antibodies made from these techniques include,among others, chimeric antibodies and humanized antibodies. A chimericantibody combines DNA encoding regions from more than one type ofspecies. For example, a chimeric antibody may derive the variable regionfrom a mouse and the constant region from a human. A humanized antibodycomes predominantly from a human, even though it contains nonhumanportions. Like a chimeric antibody, a humanized antibody may contain acompletely human constant region. But unlike a chimeric antibody, thevariable region may be partially derived from a human. The nonhuman,synthetic portions of a humanized antibody often come from CDRs inmurine antibodies. In any event, these regions are crucial to allow theantibody to recognize and bind to a specific antigen.

As noted, murine antibodies play an important role in antibodytechnology. While useful for diagnostics and short-term therapies,murine antibodies cannot be administered to people long-term withoutincreasing the risk of a deleterious immunogenic response. Thisresponse, called Human Anti-Mouse Antibody (HAMA), occurs when a humanimmune system recognizes the murine antibody as foreign and attacks it.A HAMA response can cause toxic shock or even death. Chimeric andhumanized antibodies reduce the likelihood of a HAMA response byminimizing the nonhuman portions of administered antibodies.Furthermore, chimeric and humanized antibodies have the additionalbenefit of activating secondary human immune responses, such as antibodydependent cellular cytotoxicity.

An “intact” antibody is one that comprises an antigen-binding variableregion as well as a light chain constant domain (CL) and heavy chainconstant domains, CH1, CH2 and CH3. The constant domains may be nativesequence constant domains (e.g., human native sequence constant domains)or amino acid sequence variant thereof. The intact antibody may have oneor more “effector functions” which refer to those biological activitiesattributable to the Fc region (a native sequence Fc region or amino acidsequence variant Fc region) of an antibody. Examples of antibodyeffector functions include C1q binding; complement dependentcytotoxicity; Fc receptor binding; antibody-dependent cell-mediatedcytotoxicity (ADCC); phagocytosis; down regulation of cell surfacereceptors (e.g., B cell receptor; BCR), etc.

Depending on the amino acid sequence of the constant domain of theirheavy chains, intact antibodies can be assigned to different “classes.”There are five major classes of intact antibodies: IgA, IgD, IgE, IgG,and IgM, and several of these may be further divided into “subclasses”(isotypes), e.g., IgG1, IgG2, IgG3, IgG4, IgA, and IgA2. The heavy-chainconstant domains that correspond to the different classes of antibodiesare called α, δ, ε, γ, and μ, respectively. The subunit structures andthree-dimensional configurations of different classes of immunoglobulinsare well known.

In embodiments, the first ultrafiltering concentrates the first antibodypreparation to provide the second antibody preparation having anantibody concentration of about 30 grams per liter, and the secondultrafiltering concentrates the intermediate antibody preparation(obtained from diafiltering) to provide the third antibody preparationhaving an antibody concentration of, for example, about 170 to about 200grams per liter. The first ultrafiltering and the second ultrafilteringcan be accomplished with the same ultra-filter membrane, and if desired,within the same vessel or process circuit, for example, to minimizehandling, losses, leakage, and like impacts on yield, efficiency, andeconomy. The first ultrafiltering and the second ultrafiltering can beaccomplished with any suitable ultrafilter apparatus or ultrafiltermembrane. Many suitable ultrafilter apparatus and ultrafilter membranes,which are capable of tangential flow filtration (TFF) operation toaccomplish the ultrafiltrations and diafiltration, are commerciallyavailable, such as from Millipore, Pall Corp., Sartorius, and likevendors. In embodiments, a suitable ultra-filter membrane can be, forexample, any regenerated cellulose composite, which composite has arelatively low protein adsorption profile compared to other availableultra-filter membranes, such as, polyethersulfone.

The diafiltering operation exchanges a first buffer composition presentin the first and second antibody preparations for a second bufferdesired in the third antibody preparation. In embodiments, the firstbuffer can comprise, for example, a mixture of aqueous sodium chlorideand a TRIS buffer, and the second buffer can comprise, for example, amixture of aqueous histidine chloride and arginine chloride. Thediafiltering can accomplish a buffer exchange at constant volume,constant concentration, or both. In embodiments, the diafilteringaccomplishes a buffer exchange at constant volume and constantconcentration. The diafiltering can accomplish a buffer exchange, forexample, of from about 5 to about 15 fold volumes (i.e. diavolumes). Thediafiltering can also accomplish a buffer exchange, for example, ofabout 8-fold volumes (8 diavolumes), that is, 8 times the volume of thesample containing the antibody preparation to be exchanged. For example,a 10 liter antibody preparation can be diafiltered with a 5 fold(diavolumes) or 50 liter volume of exchange buffer. The exchange volumeand preferences for exchange volumes considers a balance of factors, forexample, process through-put efficiencies, product purity, governmentaland customer-patient acceptability standards, and like standards, andcan depend on, for example, the concentration and type of buffer (e.g.,the first buffer) in the first antibody preparation, and likeconsiderations.

The first ultrafiltering, the second ultrafiltering, and thediafiltering are preferably accomplished with tangential flow filtration(TFF mode) across an ultra-filter membrane, and the ultra-filtermembrane is preferably the same membrane for each step. The yield ofproduct in the final pool (i.e., the third antibody preparation) can be,for example, greater than about 70 weight percent, such as from about 80to about 100 weight percent based on the weight of antibodies in thefirst antibody preparation. The yield of the third antibody preparationcan be, in embodiments, greater than about 90 weight %, in embodiments,greater than about 95 weight percent, and in embodiments, even greaterthan about 98 weight %, based on the weight of antibodies in the firstantibody preparation.

The first ultrafiltering can have a recirculation rate of, for example,from about 50 to 1,000 mL/min, and preferably from about 100 to 1,000mL/min. The recirculation rate can be scaled in accordance with theavailable membrane area, for example, membrane areas of 5, 20, 200,1,000 square feet, and like areas permit increasingly higherrecirculation rates. Thus, a suitable scaled recirculation rate, inembodiments, can be, for example, from about 0.5 L/min/ft² to about 5L/min/ft². The ultrafiltering and diafiltrating can be accomplished, forexample, at transmembrane pressures of from about 5 to about 50 p.s.i.The ultrafiltering and diafiltrating can be accomplished, for example,at transmembrane pressures of from about 10 to about 50 p.s.i. Inembodiments of the present disclosure there is provided a process forpreparing an antibody concentrate for a more dilute antibodyformulation, the antibody concentrate having a minimum bio-burden, forexample, of less than or under a detectable limit, such as, less thanabout 100 CFU/mL.

Antibody compositions of the disclosure can be, for example,concentrated monoclonal antibody preparation for administration tohumans, such as at a concentration of greater than or equal to about 100g/L (mg/mL), such as about 120 to about 170 g/L.

The antibody compositions of the disclosure can be, for example,immunoglobulins, such as from the group IgA, IgD, IgE, IgG, and IgM;sub-classes thereof; recombinants thereof; fragments thereof; andmixtures thereof of any of the foregoing. A preferred antibodycomposition of the disclosure includes recombinant humanized anti-IgEantibodies. The antibody compositions of the disclosure can include abuffer. A preferred buffer can be, for example, a mixture of aqueoushistidine chloride and arginine chloride.

The preparative processes of the disclosure are preferably accomplishedin the same apparatus and without operator intervention or with minimaloperator intervention, for example, as illustrated in FIG. 1.

The first antibody preparation can be provided or prepared using avariety of chemical, physical, mechanical or non-mechanical, orbiochemical methods, such as, grinding, ultrasonication, homogenization,enzymatic digestion, solvent extraction, centrifugation, chromatography,and like methods, and combinations thereof, see for example, the abovementioned R. Hatti-Kaul et al., “Downstream Processing inBiotechnology,” in Basic Biotechnology, Chap. 9. The third antibodypreparation can be further processes, if desired, using for example,nanofiltration (to remove, e.g., divalent ions), reverse osmosis (toremove, e.g., monovalent ions), and like liquid purification methods.The third antibody preparation of the present disclosure can bepackaged, stored, or directly used. The third antibody preparation canbe further processed, if desired, using for example, additionalconcentration steps, such as drying, lyophilization,lyophilization-reconstitution, and like methods. The resultingconcentrated third antibody product can be reconstituted at a latertime, if desired, with a suitable liquid.

Referring to the figures, FIG. 1 illustrates an apparatus, inembodiments of the present disclosure, for accomplishing the preparativeprocess including an ultrafiltration-difiltration system (100) having anTFF ultra-filtration-difiltration (UF-DF) unit (110), having a UF-DFmembrane (115), which is in communication with recirculation tank (120)which tank serves as a main feed and retentate reservoir. Inembodiments, tank (120) can have a temperature control systemcomprising, for example, an insulating jacket (125), a thermostatic ortemperature controlled heating element (126), such as a rheostatresistive heater element or a circulating heated liquid system whichincludes a heater (not shown), a flow regulator (127), such as arecirculating pump, and a suitable heat transfer fluid, such as eitherwater, glycols, or mixtures thereof. All in-circuit components orcomponent contributing to in-circuit flow or processing, such as pipes,valves, pumps, tanks, and like components, can be optionally insulatedor optionally adapted for external heating to maintain close controlover temperature specifications and to avoid temperature excursions inthe recirculating fluid loop within and between filter chamber (110) andrecirculation tank (120). In embodiments, for example, when the system(100) is accomplishing the first ultrafiltration or firstultrafiltering, such as in a fed-batch mode, the system can include anoptional feed tank (128) which is in fluid communication withrecirculation feed tank (120) and can be used to, for example, make-up,replenish, or supplement the depleted liquid phase from recirculationtank (120).

A pump (130) pumps feed liquid from tank (120) through the UF/DF unit(110) and thereafter recirculates the resulting retentate (thenon-filtered or membrane excluded portion of the feed liquid) torecirculating tank (120). A second tank (140) holds and optionally pumps(not shown) a buffer into the main circuit (110-120 loop) during theconstant volume diafiltration. For example, the addition rate and volumeof the buffer introduced into the main circuit is preferably at the samerate and volume at which filtrate leaves the main circuit throughmembrane (115). Buffer tank (140) can be optionally insulated withjacket (143) and can include the equivalent of the abovementionedheating element and a recirculating pump (not shown). An optional inertgas source (145), such as nitrogen, or other compressed gas sources canbe used, for example, for product recovery, to pressurize the retentatereturn, exclude oxygen, for flushing, for cleaning, for membraneintegrity testing, and like operations. A third tank (160) is used tocollect and recover the filtrate exiting the unit (110). Valves (150,170) can be used as appropriate to regulate the direction and optionallythe liquid flow rate in the system. All values and pumps can be actuatedmanually, by coordinated computer control, or both. An optional forthtank (190) and exit stream can provide an ancillary waste-flush, productrecovery, or monitoring system, for example, when equipped with anoptional monitoring device (180), such as an optical density meter,optional filter(s) (185) such as a guard filter, product filter, andlike optional subsystems. In embodiments, the main fluid circuit(110-120 loop) can optionally be equipped with an in-line monitoringsystem.

The concentrated antibody preparations prepared by processes of thepresent disclosure can be used for human therapeutic administration,including immunoglobulin products, for either intramuscular (IMIG) orintravenous (IVIG) administration. The concentrated antibodypreparations of the disclosure can include a stabilizer, for example, abuffered amino acid salt solution, simple sugars, or like stabilizers,suitable ions chelators, such as EDTA or citrate ion, and combinationsthereof, see for example, Wang, Y.-C. J. et al, “Parenteral formulationsof proteins and peptides: stability and stabilizers,” J. Parenteral Sci.Technol., 42, Suppl. S3-S26 (1988). Derwent Abstract of JP01268646A(AN89-359879) reports that the application describes an injectionpreparation of an IgG₃ monoclonal antibody having a concentration of 0.1micrograms/mL to 100 mg/mL. Subject matter disclosed in thesepublications is believed to be outside the scope of the presentdisclosure.

Preparations according to the disclosure can be substantially free fromaggregates. Acceptable levels of aggregated contaminants would be lessthan, for example, about 5 weight %, and ideally less than 2 weight %.Levels as low as 0.2 weight % can be achieved, although aggregatedcontaminants of about 1 weight % is more typical. The preparation inembodiments, can also preferably be free from excipients traditionallyused to stabilize polyclonal formulations, for example glycine and/ormaltose.

The present disclosure can provide a monoclonal antibody preparation foradministration to a human characterized in that the antibody in thepreparation is a recombinant antibody and can be at a concentration of100 mg/mL or greater, preferably greater than 150 mg/mL. The preparationis preferably substantially free from of any protein aggregation.

The pH of pharmaceutical formulations of the disclosure will depend uponthe particular route of administration. However, in order to maximizethe solubility of the antibody in the concentrated solution, the pH ofthe solution should be different from the pH of the isoelectric point(pI) of the antibody.

In embodiments of the disclosure, the monoclonal preparation can beenvisaged for use in human therapy. Various human disorders can betreated such as cancer or infectious diseases, for example, thosementioned above, and immune dysfunction such as T-cell-mediateddisorders including severe vasculitis, rheumatoid arthritis, systemiclupus, also autoimmune disorders such as multiple sclerosis, graftversus host disease, psoriasis, juvenile onset diabetes, Sjogrens'disease, thyroid disease, myasthenia gravis, transplant rejection,inflammatory bowel disease, asthma, IgE mediated disorders, and likedisorders or conditions, or combinations thereof.

The disclosure therefore provides in embodiments the use of aconcentrated monoclonal antibody preparation as described herein in themanufacture of medicament for the treatment of any of the aforementioneddisorders, and like disorders. Also provided is a method of treating ahuman being, having any such disorder, comprising administering to theindividual a therapeutically effective amount of a preparation accordingto the disclosure. The dosages of such antibody preparations will varywith the conditions being treated and the recipient of the treatment,but can be, for example, in the range of about 50 to about 2,000 mg foran adult patient preferably about 100 to about 1,000 mg administereddaily or weekly for a period between 1 and 30 days, and repeated asnecessary. The doses may be administered as single or multiple doses.

Process Description. The formulation step typically exchanges thepurified bulk drug substance, for example, resulting from ion-exchangechromatography, into the final excipient composition and concentration.There was typically no purification achieved at this step except forsmall molecule removal. The emphasis was on high yield, buffer exchange,and formulation step robustness. During formulation via TFF (tangentialflow filtration), the protein-containing feed solution was pumpedthrough the membrane system and back to the recycle (recirculation)vessel. The TFF membrane retained the protein (as part of the retentate)while the filtrate (or permeate) was driven through the membrane bypressure. The pressure is called the transmembrane pressure (TMP) and istypically controlled using a retentate pressure control valve. Theprocess was usually achieved by a sequence of a first ultrafiltering(concentration), diafiltering (constant volume buffer exchange), and asecond ultrafiltering (further concentration). The number of diavolumes(volumetric equivalents) necessary to remove process buffer componentscan be readily calculated or determined experimentally.UF/DF Process Generally for anti-IgE. The pH of an anion-exchange poolfrom chromatography was adjusted to a pH of about 6 using 0.5 M aqueousphosphoric acid. The pH adjusted anion-exchange pool was formulated byultrafiltration/diafiltration (UF/DF) process of the present disclosureusing a membrane having a nominal molecular cut off of 10,000-30,000Daltons. Prior to processing, the UF membrane was equilibrated withdiafiltration buffer (0.02 M histidine, 0.2 M arginine-HCl, pH 6).

The product from an anionic exchange (anion-exchange pool) was thenloaded on the system and was concentrated to an intermediateconcentration by the first ultrafiltering. The pool was then diafiltered(8× or diavolumes) into its formulation (0.02 M histidine, 0.2 Marginine-HCl, pH 6). The pool was then concentrated by a secondultrafiltering to a final bulk concentration of >170 g/L and recoveredthrough a 0.22 micrometer sterile filter. The entire UF/DF process wasperformed at a temperature set point of about 45 degrees C. Thistemperature control was achieved using temperature control of theincoming anion-exchange pool, the diafiltration buffer, and the use ofajacketed recirculation vessel for the UF/DF process as illustratedherein.

After UF/DF, the recovered pool was diluted (i.e., conditioned) to abulk concentration of about 150 g/L in 0.02 M histidine, 0.2 Marginine-HCl, 0.04% polysorbate-20, pH 6 (final formulation). During theconditioning steps the temperature of the bulk was allowed to return toambient temperature. After conditioning, the formulated bulk was againrecovered through a 0.22 micrometers sterile filter.

The UF/DF system can be regenerated with 0.1 N sodium hydroxide andsanitized with 1.4% Minncare®. When not in use the system can be storedin 0.1N aqueous sodium hydroxide. The UF/DF membranes can be stored, forexample, in a 0.1% Roccal®/20% glycerol-water solution betweencampaigns.

General Ultrafiltration/Diafiltration Process Procedures

Operating Parameters Feed flow rate at 0.5 L/min/ft². A constantretentate pressure (e.g., 10 psig) control was used for cleaning andpre-use equilibration, whereas C_(wall), constant retentate pressure orconstant TMP was used for processing.

Pre-Use Equilibration: The following preparations were accomplished oncleaned Pellicon-2 cassette membranes prior to use to assure themembranes were properly equilibrated.

Volume (L/ft²) Solution (room temp) Mode — — SPFO 1.0 WFI SPFO 1.0 DFbuffer SPFO 0.5 DF Buffer TRFO, 10 minutes — — SPFO

Process Use: The following was performed on the resulting initialanion-exchange pool (Q-pool) obtained from a preceding separation step,for example, a Q-Sepharose chromatography step:

a first ultrafiltering or first ultrafiltration (UF1) to a concentrationfrom about 5 g/L to a concentration for difiltration (C_(DF));

diafiltering or diafiltration (DF1) with four (4) difiltration volumes(DV) with the DF buffer;

continued diafiltering (DF2) with four (4) difiltration volumes (DV) ofDF buffer;

a second ultrafiltering or second ultrafiltration (UF2) to a finalconcentration (C_(Final)); and

optional product recovery.

The foregoing steps were typically accomplished at low dP Recycle (mix),for example, 15 min.

Post-Use Cleanout: The following tabulated sequence and conditions wereused for cleanout on the Pellicon-2 cassette membranes immediatelyfollowing use.

Volume Solution (L/ft²) (room temp) Mode 1.0 0.1N NaOH SPFO 0.5 0.1NNaOH TRFO, 30 minutes — — SPFO 1.0 WFI SPFO 0.5 300 ppm TRFO, 30Minncare ® minutes — — SPFO 1.0 WFI SPFO — — Integrity Test @ 30 psig0.5 0.1N NaOH TRFO, 15 minutes storage

DEFINITIONS FOR MODES OF OPERATION IN TFF

Single Pass with Filtrate Open (SPFO). The retentate and filtrate aredirected to drain. Filtrate valve open.

Total Recycle with Filtrate Open (TRFO). The retentate and filtrate aredirected to recycle vessel. Filtrate valve open.

Fed-Batch Ultrafiltration (FB-UF). The retentate is directed to therecycle tank, the filtrate directed to drain, and the incoming pooltransferred into the recycle tank.

Batch Ultrafiltration (B-UF). The retentate is directed to the recycletank and the filtrate is directed to drain.

Diafiltration (DF). The retentate is directed to the recycle tank, thefiltrate is directed to drain, and the diafiltration buffer istransferred into recycle tank.

dP refers to differential pressure.

Product Transfer. The ultrafilter membrane unit and recycle tank areopen to the pool tank. The nitrogen overlay pressure is controlled. Thepool is transferred first using the recycle pump and then using a manualperistaltic pump.

Feed Transfer. The incoming pool is pumped into the recycle tank.

Total Recycle with Filtrate Closed (TRFC). The retentate is directed toa recycle vessel. Filtrate valve closed.

“Q-pool” refers to the protein pool resulting from, for example, apreceding Q-Sepharose chromatography step which has been conditionedwith buffer, also referred to as the “conditioned pool.”

WFI refers to water-for-injection.

EXAMPLES

The following examples serve to more fully describe the manner of usingthe above-described disclosure, as well as to set forth the best modescontemplated for carrying out various aspects of the disclosure. It isunderstood that these examples in no way serve to limit the true scopeof this disclosure, but rather are presented for illustrative purposes.

Example 1

High Concentration Formulation of rhuMAb E25 A pilot scale UF system wasused to concentrate/formulate rhuMAb E25 (a recombinant human monoclonalantibody that targets IgE, U.S. Pat. No. 6,172,213). A Millipore PeliconUltrafiltration/Diafiltration system was assembled with a 5.7-sqft,10,000-dalton regenerated cellulose composite membrane. The systemconsisted of a membrane holder, a Waukeskaw Model 6 rotary lobe feedpump, ½″ 316 L stainless steel recirculation piping, and a recirculationvessel. Pressure indicators/transmitters (Anderson) were located at theinlet (FEED), outlet (RETENTATE) and permeate (FILTRATE) of the membraneholder. Flow meters (Yokogawa ADMAG) were located at the inlet (FEED)and permeate (FILTRATE) of the membrane holder. A back-pressureregulating valve (Mikroseal) was located at the outlet of the membraneholder to control the retentate pressure and effect the transmembranepressure (TMP). A 40-liter 316 L stainless steel jacketed tank was usedfor the recirculation vessel. This tank was fitted with a levelindicator, top-mounted agitator (Lightnin), vortex breaker and bottomvalve (NovAseptic). Temperature control was achieved through the use oftemperature modulated glycol fed to the jacket of the tank.

During this run, the feed flow rate was set to a constant rate of 2.85L/min (0.5 L/min/ft²). During all pre-use and post-use operations theretentate pressure control was set to a constant of 10 psig. During theultrafiltering and diafiltering operations the system used a C_(wall)control scheme to control the flux through the membrane, see for exampleR. van Reis, et al., Constant C_(wall) Ultrafiltration Process Control,J. of Membrane Science, 130 (1997), 123-140.

Prior to the process, the system storage solution (0.1N NaOH) wasflushed in a single pass to drain mode, first with 2 L/ft² purifiedwater (PW) and then 1 L/ft² diafiltration buffer (50 mM Histidine/pH6.0). After the flushes, the system was equilibrated by recirculating0.5 L/ft² diafiltration buffer for 10 min. The pH of the recirculatedsolution was checked to confirm the equilibration. The level in the tankwas then reduced to a minimum measurable value to minimize dilution ofthe incoming protein pool. The protein pool resulting from a precedingQ-Sepharose chromatography step was measured to be 3.2 g E25/L and had avolume of 43.1 L. The protein was in a solution of 25 mM TRIS buffer andabout 200 mM NaCl and pH adjusted to 6.2. To begin the run the proteinpool was transferred to the recirculation vessel. In the vessel the poolwas agitated via the top mounted impeller and the temperature wasmaintained at ambient (20-25° C.).

During the process the pool was concentrated in UF1 mode to 50 g E25/L(about 2.8 L). At the beginning of diafiltration the temperature setpoint of the recirculation vessel was increased to 40° C. The increasein temperature and control was affected by flowing warm glycol throughthe outer jacket of the tank. The pool was then diafiltered with 8diavolumes of diafiltration buffer. The diafiltration was performed at aconstant volume, which was achieved by matching the flow rate ofbuffered solution being transferred into the recirculation tank to theflow rate of filtrate being removed from the system. At the end of thediafiltration, the pool was further concentrated in UF2 mode. This phasewas also performed using an elevated temperature set point of 40° C. Thetarget for this final concentration was 110 g/L. This was achievedwithout the need to reduce the feed flow rate. Next, a low pressure dropmixing was performed where the feed pump was controlled to maintain a5-10 psig pressure drop across the feed channel. A sample was pulledfrom the recirculation tank and a final bulk concentration ofapproximately 120 g/L was measured. Table 1 summarizes the throughputand flux results of UF1, DF (DF1+DF2), and UF2.

TABLE 1 Normalized Throughput Normalized Flux Process Phase (g/ft²/hr)(LMH/psig) UF1 13.8 4.97 DF 13.8 2.92 UF2 181.4 1.64

FIG. 2 shows the observed or measured process values over time for thefeed flow rate (210), tank temperature (220), fed dP (230), TMP (240),and filtrate flow rate (250) parameters during the various phases ormode of the process including UF1 (10), DF (20), UF2 (30).

FIG. 3 shows the observed or measured process values over time for theE25 concentration (310), flux (320), and TMP (240).

FIG. 4 shows the observed or measured process values over time forpressure drop versus protein concentration observed for UF1 (410) andUF2 (420) at 37° C.

The protein pool was recovered by a series of steps. First the pool inthe recirculation tank was pumped from the tank through a Millipac 200,0.22 microns sterilizing grade filter using the rotary lobe feed pump.Next the protein solution was displaced from the piping and membraneunit with a 5 psig nitrogen gas blow down applied to the highest pointon the retentate line. The final phase was a blow down of the tank andfeed line, also using the 5 psig nitrogen gas.

The product recovery was believed to be improved compared to Example 1when conducted at ambient temperature because the elevated temperatureused in one or more of the ultrafiltering, diafiltering, or recoverysteps reduced viscous effects. For example, when the temperature controlwas turned off during product recovery, the system slowly cooled duringthis operation causing difficulties for recovery from the membrane unit.Alternatively, the recovery can be performed first from the membraneholder and then from the recirculation vessel.

To determine the mass of loss during recovery, 1.74 L of DF buffer wasadded to the system and recirculated for about 5 minutes and recoveredusing the same sequence as described above. This volume was thenanalyzed for protein concentration with the other pools. Table 2summarizes the results.

TABLE 2 Concentration Yield or Volume (L) (g/L) Mass (g) {Loss} (%)Q-Pool 43.1 3.2 137.9 100 Recovered 0.99 120 118.8 86.1 Pool Buffer 1.749.8 17.1 12.4 Flush Filtrate 65.3 0.04 2.6 1.9

Post processing, the membrane was regenerated using 0.1N NaOH, 1 L/ft²single pass flush followed by 0.5 L/ft² total recirculation for 30 min.This was followed by 1 L/ft² PW (pure water) flush. This was followed bya total recirculation of 300 ppm Mirncare® solution for 30 min. Thesystem was again flushed with 1 L/ft² PW and finally recirculated for 15min with 0.1N NaOH and stored. The recovered pool was diluted to 80 gE25/L and conditioned into the final formulation of 50 mM histidine/150mM trehalose/0.02% polysorbate 20/pH 6.0. Product quality was assessedby size exclusion chromatography (SEC) for both the incoming Q-Pool andfinal recovered bulk. This data is summarized in Table 3.

TABLE 3 SEC Results (% Pool monomer) Q-Pool 99.8 Final Bulk 99.8

Comparative Example 2

High Concentration Formulation of rhuMAb E25 at Ambient TemperatureExample 1 was accomplished with the following exceptions. Prior to theprocess, the system storage solution (0.1N NaOH) was flushed in a singlepass to drain mode first with 2 L/ft² purified water (PW) and then 1L/ft² diafiltration buffer (20 mM histidine/pH 6.0). After the flushesthe system was equilibrated by recirculating 0.5 L/ft² diafiltrationbuffer for 10 min. The pH of the recirculated solution was checked toconfirm the equilibration. The level in the tank was then reduced to aminimum measurable value to minimize dilution of the incoming proteinpool.

The protein pool resulting from the preceding Q-Sepharose chromatographystep was measured to be 3.3 g E25/L and had a volume of 33.3 L. Theprotein was in a solution of 25 mM TRIS buffer and about 200 mM NaCl andpH adjusted to 6.2. To begin the run the protein pool was transferred tothe recirculation vessel. In the vessel the pool was agitated via thetop mounted impeller and the temperature was maintained at ambient(20-25° C.). During the process the pool was concentrated down in UF1mode to 50 g E25/L (about 2.2 L). The pool was then diafiltered with 8diavolumes of diafiltration buffer. The diafiltration was performed at aconstant volume, which volume was achieved by matching the flow rate ofbuffered being transferred into the recirculation tank to the flow rateof filtrate being removed from the system. The diafiltration was alsoperformed at ambient temperature. At the end of the diafiltration, thepool was further concentrated in UF2 mode. The target for this finalconcentration was 110 g/L. However, due to a high pressure drop acrossthe feed channel, this concentration was not achieved. In an attempt toachieve this concentration the feed flow rate was reduced to 1.4 L/minat a bulk concentration of about 80 g E25/L because the pressure dropacross the feed channel had reached 50 psig. UF2 was continued until ahigh pressure drop of 50 psig again was reached and the process wasstopped. Next, a low pressure drop mixing was attempted where the feedpump was used to maintain a 5 psig pressure drop across the feedchannel. Again, the viscous nature of the protein solution made itdifficult to achieve since the rotary lobe pump reached excesspressures. A sample was pulled from the recirculation tank and a finalbulk concentration of approximately 104 g/L was measured. Table 4summarizes throughput and flux measured during the UF1, DF (DF1+DF2),and UF2 phases.

TABLE 4 Normalized Process Throughout Normalized Flux Phase (g/ft²/hr)(LMH/psig) UF1 14.5 5.31 DF 9.5 1.47 UF2 144.6 0.78

FIG. 5 shows the observed or measured process values over time for thefeed flow rate (210), tank temperature (220), fed dP (230), TMP (240),and filtrate flow rate (250) parameters during the various phases ormode of the process including UF1 (10), DF (20), UF2 (30).

FIG. 6 shows the observed or measured process values over time for theE25 concentration (310), flux (320), and TMP (240).

FIG. 7 shows the observed or measured process values over time forpressure drop versus protein concentration observed for UF1 (410) andUF2 (420) at 24° C.

The protein pool was recovered in steps. First, the pool in therecirculation tank was pumped from the tank through a Millipac 200, 0.22microns sterilizing grade filter using the rotary lobe feed pump. Nextthe protein solution was displaced from the piping and membrane unitwith a 5 psig nitrogen gas blow down applied to the highest point on theretentate line. The product recovery from this was very poor due to theviscous nature of the solution. The final phase was a blow down of thetank and feed line, also using the 5 psig nitrogen gas.

To determine the mass of loss during recovery, 1.85 L of DF buffer wasadded to the system and recirculated for about 5 minutes and recoveredusing the sequence of Example 1. This volume was then analyzed forprotein concentration with the other pools. Table 5 summarizes theresults.

TABLE 5 Concentration Yield or Volume (L) (g/L) Mass (g) {Loss} (%)Q-Pool 33.3 3.3 109.9 100 Recovered 0.77 104.4 80.4 73.1 Pool Buffer1.85 14.7 27.2 24.7 Flush Filtrate 52.2 0.03 1.6 1.5

Post process, the membrane was regenerated using 0.1N NaOH, 1 L/ft²single pass flush followed by 0.5 L/ft² total recirculation for 30 min.This was followed by 1 L/ft2 PW flush. This was followed by a totalrecirculation of 300 ppm Minncare® solution for 30 min. The system wasagain flushed with 1 L/ft² PW and finally recirculated for 15 min with0.1N NaOH and stored. The recovered pool was diluted to 80 g E25/L andconditioned into the final formulation of 20 mM histidine/250 mMsucrose/0.02% polysorbate 20/pH 6.0. Product quality was assessed bysize exclusion chromatography (SEC) for both the incoming Q-Pool andfinal recovered bulk. This data is summarized in Table 6.

TABLE 6 SEC Results (% Pool monomer) Q-Pool 99.8 Final 99.8 Bulk

Example 3

High Concentration Formulation of rhuMAb E26 with Initial Fed-Batch ModeExample 1 was repeated with the following exceptions. Theconcentrate/formula was rhuMAb E26 (a recombinant human monoclonalantibody that targets IgE). The products from this example were used intoxicology assessment. The Millipore PeliconUltrafiltration/Diafiltration system was assembled with a 11.4-sqft30,000-Dalton regenerated cellulose composite membrane. The feed flowrate was set to a constant rate of 5.0 L/min (0.44 L/min/ft²). Duringthe ultrafiltration and diafiltration operations the retentate pressurewas maintained between about 6-8 psig. The protein pool resulting fromthe preceding Q-Sepharose chromatography step was measured to be 6.7 gE26/L and had a volume of 59.3 L.

Because the incoming pool was larger then the recirculation vessel, theUF1 process began in fed-batch mode. In this mode, Q-Pool was added tothe recirculation vessel at approximately the same rate as filtratepasses through the TFF membrane to drain. After the remaining Q-Pool hadtransferred to the recirculation vessel, the UF1 process continued inbatch mode. During the UF1 the pool was concentrated to 50 g E26/L(about 7.9 L). At the beginring of diafiltration the temperature setpoint of the recirculation vessel was increased to 40° C. The increasein temperature and control was affected by flowing warm glycol throughthe outer jacket of the tank. The pool was then diafiltered with 8diavolumes of diafiltration buffer. The diafiltration was performed at aconstant volume which was achieved by matching the flow rate of bufferedbeing transferred into the recirculation tank to the flow rate offiltrate being removed from the system. At the end of the diafiltration,the pool was further concentrated in UF2 mode to a final concentrationof 109 g E26/L (3.6 L). This phase was also performed using an elevatedtemperature set point of 40° C. Next a low pressure drop mixing wasperformed where the feed pump was controlled to maintain a 5-10 psigpressure drop across the feed channel. Table 7 summarizes the throughputand flux results of UF1, DF (DF1+DF2), and UF2.

TABLE 7 Normalized Process throughput Normalized Flux Phase (g/ft²/hr)(LMH/psig) UF1 26.1 3.71 DF 19.2 2.34 UF2 174.2 1.80

FIG. 8 shows the observed or measured process values over time for thefeed flow rate (210), tank temperature (220), fed dP (230), TMP (240),and filtrate flow rate (250).

FIG. 9 shows the observed or measured process values over time for theE26 concentration (910), flux (920), and TMP (940).

FIG. 10 shows the observed or measured process values over time forpressure drop versus protein concentration observed for UF1 (1010) andUF2 (1020).

Just prior to product recovery, a 10 mL sample was analyzed fordetection and a titer of bioburden. A typical reject limit is 1,000Colony Forming Units (CFU) per mL. The results of this test were 1.8CFU/mL, a suitable value at this step and well below the reject limit.To determine the mass of loss during recovery, 908.1 mL of DF buffer wasadded to the system and recirculated for about 5 minutes and recoveredusing the same sequence described above. This volume was then analyzedfor protein concentration with the other pools. Table 8 summarizes theresults.

TABLE 8 Concentration Yield or Volume (L) (g/L) Mass (g) {Loss} (%)Q-Pool 59.3 6.7 397.3 100 Recovered 3.41 109.1 372.0 93.6 Pool Buffer0.908 20.4 18.5 4.7 Flush Filtrate 120 n/d n/d n/d

The recovered pool was diluted to 80 g E26/L and conditioned into thefinal formulation of 50 mM histidine/150 mm trehalose/0.02% polysorbate20/pH 6.0. Product quality was assessed by size exclusion chromatography(SEC) for the incoming Q-Pool, the retentate pool after UF1, theretentate pool after DF, and final recovered bulk. This data issummarized in Table 9.

TABLE 9 SEC Results Pool (% monomer) Q-Pool 99.8 End of 99.8 UF1 End of99.8 DF Final 99.8 Bulk

Example 4

High Concentration Formulation of rhuMAb E26 for ToxicologyEvaluation-Comparison of 10 kD and 30 kD Example 3 was repeated with thefollowing exceptions. Two pilot scale UF systems were used toconcentrate/formulate rhuMAb E26. Two Millipore PeliconUltrafiltration/Diafiltration systems were assembled with a 11.4-sqft,regenerated cellulose composite membrane, one with 10,000-Dalton poresize and the other a 30,000-Dalton pore size. The retentate pressureswere maintained at about 6-9 psig.

10 kD Process

The protein pool resulting from the preceding Q-Sepharose chromatographystep was measured to be 5.85 g E26/L and had a volume of 62.4 L. Duringthe UF1, the pool was concentrated to 50 g E26/L (about 7.3 L). At theend of the diafiltration, the pool was further concentrated in UF2 modeto a final concentration of 107.5 g E26/L (3.4 L). Table 10 summarizesthe throughput and flux results of UF1, DF, and UF2.

TABLE 10 Process Normalized Normalized Flux Phase throughput (g/ft²/hr)(LMH/psig) UF1 21.8 3.6 DF 15.9 2.6 UF2 137.4 1.93

To determine the mass of loss during recovery, 987 mL of DF buffer wasadded to the system and recirculated for about 5 minutes and recoveredusing the same sequence described above. This volume was then analyzedfor protein concentration with the other pools. Table 11 summarizes theresults.

TABLE 11 Concentration Yield or Volume (L) (g/L) Mass (g) {Loss} (%)Q-Pool 62.4 5.85 365.4 100 Recovered 3.38 107.5 361.7 98.9 Pool Buffer0.987 19.9 19.6 5.4 Flush Filtrate 125 n/d n/d n/d

FIG. 11 shows the observed or measured process values over time for thefeed flow rate (210), tank temperature (220), fed dP (230), TMP (240),and filtrate flow rate (250) over the various phases or mode of theprocess including UF1 (10), DF (20), UF2 (30), and low dP (40), for the10 kD process.

FIG. 12 shows the observed or measured process values over time for theE26 concentration (1210), flux (1220), and TMP (1240) over the variousphases or mode of the process including UF1 (10), DF (20), UF2 (30), andlow dP (40), for the 10 kD process.

FIG. 13 shows the observed or measured process values over time forpressure drop versus protein concentration observed for UF1 (1310) andUF2 (1320) for the 10 kD process.

30 kD Process

The protein pool resulting from the preceding Q-Sepharose chromatographystep was measured to be 5.85 g E26/L and had a volume of 64.5 L. Duringthe UF1 the initial pool was concentrated to 50 g E26/L (about 7.5 L).At the end of the diafiltration, the pool was further concentrated inUF2 mode to a final concentration of 117.5 g E26/L (3.2 L). Table 12summarizes the throughput and flux results of UF1, DF, and UF2.

TABLE 12 Normalized Process throughput Normalized Flux Phase (g/ft²/hr)(LMH/psig) UF1 25.5 4.01 DF 17.6 2.39 UF2 180.5 1.57

To determine the mass of loss during recovery, 918 mL of DF buffer wasadded to the system and recirculated for about 5 minutes and recoveredusing the same sequence described above. The recovered pool was dilutedto 80 g E26/L and conditioned into the final formulation of 50 mMhistidine/150 mM trehalose/0.02% polysorbate 20/pH 6.0. Table 13summarizes the results.

TABLE 13 Concentration Yield or Volume (L) (g/L) Mass (g) {Loss} (%)Q-Pool 64.5 5.85 377.3 100 Recovered 3.20 117.5 376.0 99.6 Pool Buffer0.918 22.7 20.8 5.5 Flush Filtrate 125 n/d n/d n/d

FIG. 14 shows the observed or measured process values over time for thefeed flow rate (210), tank temperature (220), fed dP (230), TMP (240),and filtrate flow rate (250) over the various phases or mode of theprocess including UF1 (10), DF (20), UF2 (30), and low dP (40), for the30 kD process.

FIG. 15 shows the observed or measured process values over time for theE26 concentration (1510), flux (1520), and TMP (1540) over the variousphases or mode of the process including UF1 (10), DF (20), UF2 (30), andlow dP (40), for the 30 kD process.

FIG. 16 shows the observed or measured process values over time forpressure drop versus protein concentration observed for UF1 (1610) andUF2 (1620) for the 30 kD process.

Example 5

Liquid rhuMAb E25 Scale Up Example 1 was repeated with the followingexceptions. A production scale UF system was used toconcentrate/formulate a liquid rhuMAb E25 (a recombinant humanmonoclonal antibody that targets IgE). The product can be used intherapeutic application and human bio-equivalency trials. The MilliporePelicon Ultrafiltration/Diafiltration systems were assembled with a226-sqft regenerated cellulose composite membrane, with a pore size of30,000-Dalton. Each system consisted of a membrane holder, a Viking S3Srotary lobe feed pump, 1½″ 316 L stainless steel recirculation piping,and a 250-L recirculation vessel.

One 250-liter 316 L stainless steel jacketed tank was used for therecirculation vessel. Temperature control to this tank was achieved witha temperature modulated glycol fed to the tank's jacket. The temperatureof the glycol fed to the tank jacket was raised or lowered using eithersteam-fed heat exchanger or cold glycol supply respectively.

For this run, the feed flow rate was set to a constant rate of 114 L/min(0.5 L/min/ft²). The diafiltration buffer (20 mM histidine/200 mMarginine chloride/pH 6.0) was prepared in a separate tank. Thetemperature of this buffer was set to 45° C. prior to the process. Thisenabled accurate temperature control throughout the process.

Prior to processing, the system storage solution (0.1N NaOH) was flushedin a single pass to drain mode first with 1 L/ft² water for injection(WFI) and then 1 L/ft² diafiltration buffer. After the flushes, thesystem was equilibrated by recirculating 0.5 L/ft² diafiltration bufferfor 10 min. The pH of the recirculated solution was checked to confirmthe equilibration.

The protein pool resulting from the preceding Q-Sepharose chromatographystep was measured to be 5.2562 g E25/L and had a volume of 1,141 L. Theprotein was in a solution of 25 mM TRIS buffer and about 200 mM NaCl andthe pH was adjusted to 6.2. Just prior to the run, the temperature setpoint of this pool was set to 45° C. To begin the run the protein poolwas transferred to the recirculation vessel, through a 0.22 micronssterilizing grade filter to a level of about 200 L in the tank. In thevessel the pool was agitated via a top mounted impeller and thetemperature was maintained at about (40-50° C.). Because the incomingpool was larger then the recirculation vessel, the UF1 process began infed-batch mode. In this mode, Q-Pool was added to the recirculationvessel at approximately the same rate as filtrate passes through the TFFmembrane to drain. After the remaining Q-Pool was transferred to therecirculation vessel, the UF2 process was continued in batch mode.During the UF1 mode the pool was concentrated to about 30 g E25/L (about200 L). The pool was then diafiltered with about 8 diavolumes ofdiafiltration buffer. During diafiltration the temperature wasmaintained between 40° and 50° C. The diafiltration was performed at aconstant volume, which was achieved by matching the flow rate of bufferbeing transferred into the recirculation tank to the flow rate offiltrate being removed from the system. At the end of the diafiltration,the pool was further concentrated in UF2 mode to a final concentrationset point of >170 g E25/L (35 L). This UF2 mode phase was also performedat an elevated temperature set point of 45° C.+/−5° C. Next, a lowpressure drop mixing was performed where the feed pump was controlled tomaintain a 5-10 psig pressure drop across the feed channel. A sample waspulled and a spec scan was performed to confirm the concentration priorto recovery. The concentration of this sample was 219 g E25/L. Table 14summarizes throughput and flux measured during the UF1, DF (DF1+DF2),and UF2 phases.

TABLE 14 Normalized Process throughput Normalized Flux Phase (g/ft²/hr)(LMH/psig) UF1 43.8 3.34 DF 25.9 2.46 UF2 78.9 0.66

Just prior to product recovery, a 30 mL sample was pulled and submittedfor detection and titer of bioburden. The result was <0.13 CFU/mL. Theprotein pool was recovered by a series of steps. First, the product wasdisplaced from the membrane in a single pass mode using 5 L of DF bufferadded to the retentate line. The product was filtered into a recoverytank through a 7.4 ft², 0.22 microns sterilizing-grade guard filterfollowed by a 2 ft², 0.22 microns sterilizing-grade final filter. Thepool in the recirculation tank was then pumped from the tank using therotary lobe feed pump. Next the residual protein solution was displacedfrom tank and feed line with a 5 psig nitrogen gas blow down. The finalphase was a blow down of the membrane unit, which now contained mostlyDF buffer from the initial product displacement. This phase also usedthe 5 psig nitrogen gas applied to the highest point on the retentateline. The recovered pool was diluted first to about 153 g E25/L using DFbuffer. Finally, the pool was conditioned into the final formulation of20 mM histidine/200 mM arginine-HCl/0.04% polysorbate 20/pH 6.0. Thevolumes of the recovered pool, diluted pool, and conditioned pool(Q-pool) were then each analyzed for protein concentration. Table 15summarizes the results.

TABLE 15 Volume Concentration Yield or (L) (g/L) Mass (g) {Loss} (%)Q-Pool 1,141 5.2562 5,997.3 100 Recovered 35.0 170.0 5,950.0 99.2 PoolDiluted 39.0 147.0 5,726 95.5 Pool

FIG. 17 shows feed flow rate (210), tank temperature (220), fed dP(230), TMP (240), and filtrate flow rate (250) parameters during thevarious phases or mode of the process including UF1 (10), DF1 (20), DF2(25), UF2 (30), and low dP (50).

Example 6

Liquid rhuMAb E25 Preparation Example 5 was repeated with followingexceptions. A production scale UF system was used toconcentrate/formulate liquid rhuMAb E25 (E25, a recombinant humanmonoclonal antibody that targets IgE). The Millipore PeliconUltrafiltration/Diafiltration systems were assembled with a 226-sqftregenerated cellulose composite membrane, with a pore size of30,000-dalton. Each system consisted of a membrane holder, a Viking S3Srotary lobe feed pump, 1½″ 316 L stainless steel recirculation piping,and a 250-L recirculation vessel. One 250-liter 316 L stainless steeljacketed tank was used for the recirculation vessel. The feed flow ratewas set to a constant rate of 114 L/min (0.5 L/min/ft²). During allpre-use and post-use operations the retentate pressure control was setto a constant of 10 psig. During the ultrafiltration and diafiltrationoperations the system used the C_(wall) control scheme to control theflux through the membrane. The diafiltration buffer (20 mM Histidine/200mM arginine chloride/pH 6.0) was prepared in a separate tank. Thetemperature of this buffer was set to 45° C. prior to the process. Thisenabled accurate temperature control through the entire process. Theprotein pool resulting from the preceding Q-Sepharose chromatographystep was measured to be 5.5438 g E25/L and had a volume of 1,082 L. Theprotein was in a solution of 25 mM TRIS buffer and about 200 mM NaCl andpH adjusted to 6.2. Just prior to the run, the temperature set-point ofthis pool was set to 45° C. To begin the run the protein pool wastransferred to the recirculation vessel, through a 0.22 micronssterilizing grade filter to a level of about 200 L in the tank. In thevessel the pool was agitated via the top mounted impeller and thetemperature was maintained at ambient (40-50° C.). Because the incomingpool was larger then the recirculation vessel, the UF1 process began infed-batch mode. In this mode, Q-Pool was added to the recirculationvessel at approximately the same rate as filtrate passes thought the TFFmembrane to drain. After the remaining Q-Pool had transferred to therecirculation vessel, the UF1 process continued in batch mode. Duringthe UF1 the pool was concentrated to about 30 g E25/L (about 200 L). Thepool was then diafiltered with 8 diavolumes of diafiltration buffer.During diafiltration the temperature was maintained between 40 and 50°C. The diafiltration was performed at a constant volume, which wasachieved by matching the flow rate of buffered being transferred intothe recirculation tank to the flow rate of filtrate being removed fromthe system. At the end of the diafiltration, the pool was furtherconcentrated in UF2 mode to a final concentration set-point of greaterthan 170 g E25/L (35 L). This phase was also performed at an elevatedtemperature set point of 45° C.+/−5° C. Next, a low pressure drop mixingwas performed where the feed pump was controlled to maintain a 5-10 psigpressure drop across the feed channel. A sample was pulled and a specscan was performed to confirm the concentration prior to recovery. Theconcentration of this sample was 191 g E25/L and the pool volume was31.9 L. A graph of the process parameters over time were comparable tothose observed and summarized for the above FIG. 17.

TABLE 14 Normalized throughput Normalized Flux Process Phase (g/ft²/hr)(LMH/psig) UF1 45.1 3.21 DF 25.9 2.51 UF2 121.4 0.79

Just prior to product recovery, a 30 mL sample was pulled and analyzedfor a titer of bioburden. The results of this test were below thedetection limit (<0.13 CFU/mL). The protein pool was recovered by aseries of steps. First the product was displaced from the membrane in asingle pass mode using 5 L of DF buffer added to the retentate line. Theproduct was filtered into a recovery tank through a 7.4 ft², 0.22microns sterilizing-grade guard filter followed by a 2 ft², 0.22 micronssterilizing-grade final filter. The pool in the recirculation tank wasthen pumped from the tank using the rotary lobe feed pump. Next, theresidual protein solution was displaced from tank and feed line with a 5psig nitrogen gas blow down. The final phase was a blow down of themembrane unit, which contained mostly DF buffer from the initial productdisplacement. This phase also used 5 psig nitrogen gas applied to thehighest point on the retentate line. The recovered pool was dilutedfirst to about 153 g E25/L using DF buffer. Finally the pool wasconditioned into the final formulation of 20 mM histidine/200 mMarginine-HCl/0.04% polysorbate 20/pH 6.0. The volumes of the recoveredpool, diluted pool, and conditioned pool were then analyzed for proteinconcentration. Table 15 summarizes the results. Post process, themembrane was regenerated as described above.

TABLE 15 Volume Concentration Yield or (L) (g/L) Mass (g) {Loss} (%)Q-Pool 1,082 5.5438 5,998.4 100 Recovered 34.95 167.08 5,839.8 97.4 PoolDiluted 38.2 152.14 5,810.3 96.7 Pool

Example 7

Effect of Elevated Temperature on Product Quality E25 samples at 30 g/Land 150 g/L in histidine and Q buffers were kept a various temperaturesfor 24 hours. Samples were taken for turbidity measurements and SECassays. The results of turbidity versus temperature for E25 at 30 g/L inQ buffer are shown in FIG. 18. FIG. 19 shows the amount of solubleaggregate of E25 (150 g/L in 50 mM histidine buffer, pH 6.0) observedover time and at temperatures of 23° C., 40° C., 50° C., 60° C. and 70°C. The four time intervals (time of 0 hours, 4 hours, 7.5 hours, and 24hours) for each of these temperatures is shown as the cluster of fourbars from left to right as 1810 and 1910, in FIGS. 18 and 19. Thesolution turbidity was essentially unchanged after 24 hours at 60° C. Nosignificant soluble aggregate of E25 was observed below 70° C.suggesting the product samples were substantially stable up to at least60° C. and at least 24 hours.

Example 8

Effect of Elevated Temperature on Bioburden E25 samples at 30 g/L inboth arginine and histidine buffers were inoculated with 10³ colonyforming units per mL for two challenge organisms: a Gram positive strain(Staphylococcus aureus); and one Gram negative strain (Pseudomonaschlororaphis). Samples were taken after 1.5 hours and 6 hours. Theresults shown in the bar charts of FIGS. 20 and 21 indicate that thesechallenge organisms both decreased with increasing temperature. Thethree temperature intervals (temp of 25° C., 40° C., and 50° C. hours)for each observed time interval is shown as the cluster of three barsfrom left to right as 2010 and 2110, in FIGS. 20 and 21. Theinoculations shown were conducted in arginine buffer with proteinconcentrations of 30 g/L.

Example 9

Effect of Elevated Temperature on Process Flux E25 samples at 10 g/L in0.2M arginine, 25 mM histidine, pH 6.0 buffer were evaluated for theirinfluence on flux versus transmembrane pressure (TMP). FIG. 22 showsthat raising the system temperature also increased the process fluxduring the UF/DF operations. Flux excursions at various bulkconcentrations and three different temperatures of 23° C. (2210), 40° C.(2220), and 46° C. (2230) were performed. The mass transfer coefficientand filtrate flux increased by about 2 to about 3 fold providingconsiderably reduced process times.

Example 10

High Concentration Formulation of rhuMAB anti-CD20 (“2H7”) A pilot scaleUF system was used to concentrate and formulate rhuMAb anti-CD20 (2H7; arecombinant human monoclonal antibody). Example 1 was repeated with thefollowing exceptions. The Millipore PeliconUltrafiltration/Diafiltration systems were assembled with a 17.5-sqft,regenerated cellulose composite membrane, with a pore size of30,000-Dalton. The system consisted of a membrane holder, a Viking S1 Lrotary lobe feed pump, ½″ 316 L stainless steel recirculation piping,and a 40-L recirculation vessel. Backpressure regulating valves were H.D. Baumann, Inc. The temperature of the glycol fed the tank jacket wasregulated higher or lower as needed using an electric heat exchanger, acold glycol supply, or both.

During this run, the feed flow rate was set to a constant rate of 8.5L/min (approximately 0.5 L/min/ft²). FIG. 23 displays the value trendsover time for feed flow rate (210) scaled from 0 to 20, pH (212) scaledfrom 2 to 12, filtrate flow rate (250) scaled from 0 to 5, recycle tanklevel (2320) scaled from 0 to 45, and retentate dP (2350) scaled from 0to 100 during the various phases or mode of the process including UF1(10), DF1 (20), and UF2 (30).

During the ultrafiltration and diafiltration operations the system usedconstant retentate pressure followed by a constant feed/retentate deltapressure control scheme to control the flux through the membrane. Thediafiltration buffer (30 mM sodium acetate/pH 4.9) was prepared in aseparate tank. The temperature of this buffer was set to 45° C. prior tothe process for accurate temperature control through the entire process.Prior to processing, the system storage solution (0.1 N NaOH) wasflushed in a single pass to drain mode first with 1 L/ft² water forinjection (WFI) and then 1 L/ft² diafiltration buffer. After the flushesthe system was equilibrated by recirculating 0.5 L/ft² diafiltrationbuffer for 10 min. The pH of the recirculated solution was checked toconfirm the equilibration.

The protein pool resulting from a preceding Q-Sepharose chromatographystep was measured to be 2.31 g 2H7/L and had a volume of 356 L. Theprotein was in a solution of 6 mM HEPES free acid/19 mM HEPES sodiumsalt and 25 mM sodium acetate that had been pH adjusted to 5.3 with 0.5M acetic acid. Just prior to the run, the temperature set point of thispool was set to 45° C. To begin the run the protein pool was transferredto the recirculation vessel through a 0.22 microns sterilizing gradefilter to a level of about 40 L in the tank. In the vessel the pool wasagitated via the top mounted impeller and the temperature was maintainedat 40-50° C.

Because the incoming pool was larger then the recirculation vessel, theUF1 process began in fed-batch mode (see FIG. 23). In this mode, Q-Poolwas added to the recirculation vessel at approximately the same rate atas filtrate passes thought the TFF membrane to drain. After theremaining Q-Pool has transferred to the recirculation vessel, the UF1process continued in batch mode. During the UF1 the pool wasconcentrated to about 50 g 2H7/L (about 16 L). The pool was thendiafiltered with 10 diavolumes of diafiltration buffer. Duringdiafiltration the temperature was maintained between 40 and 50° C. Thediafiltration was performed at a constant volume, which was achieved bymatching the flow rate of buffered being transferred into therecirculation tank to the flow rate of filtrate being removed from thesystem. At the end of the diafiltration, the pool was furtherconcentrated in UF2 mode to a final concentration target set point of190 g 2H7/L (4.3 L). See in FIG. 23 the incorporation of constant dPcontrol at 50 psig at the end of this phase. This phase was alsoperformed at an elevated temperature set point of 45° C.+/−5° C. Next, alow pressure drop mixing was performed where the feed pump wascontrolled to maintain a 20 psig pressure drop across the feed channel.A sample was pulled and a density measurement was performed to confirmthe concentration prior to recovery. The concentration of this samplewas 189 g 2H7/L. Table 16 summarizes the throughput and flux results.

TABLE 16 Normalized throughput Normalized Flux Process Phase (g/ft²/hr)(LMH/psig) UF1 32 4.8 DF 56 2.4 UF2 267 1.6

The protein pool was recovered by a series of steps. First the productis displaced from the membrane in a single pass mode using 0.2 L of DFbuffer added to the retentate line. The product is filtered into arecovery tank through a 0.22 microns sterilizing-grade final filter. Thepool in the recirculation tank was then pumped from the tank using therotary lobe feed pump. Next, the residual protein solution is displacedfrom tank and feed line with a 5 psig nitrogen gas blow down. The finalphase was a blow down of membrane unit, which now contains DF bufferfrom the initial product displacement. This phase also used the 5 psignitrogen gas applied to the highest point on the retentate line.

If necessary, the recovered pool was diluted first to about 175 g 2H7/Lusing dilution buffer (30 mM sodium acetate, pH 5.3). Finally, the poolis diluted down to a target concentration of 150 g 2H7/L and conditionedinto the final formulation of 30 mM sodium acetate, 7% trehalose, 0.03%polysorbate 20, pH 5, via a 7× conditioning buffer (30 mM sodiumacetate, 49% trehalose, 0.21% polysorbate 20, pH 5.3). The volumes ofthe recovered pool, diluted pool, and conditioned pool were thenanalyzed for protein concentration. Table 17 presents the results.

TABLE 17 Concentration Yield or Volume (L) (g/L) Mass (g) {Loss} (%)Q-Pool 355.81 2.31 821.92 100.0 Recovered 4.64 180.02 835.3 101.6 PoolFinal Pool 4.871 149.40 727.7 88.5 Note: Yields include loss due tosampling. Recovered pool volume and concentration include addition ofbuffer displacement.

Post process, the membrane was regenerated using 0.1 N NaOH, 1 L/ft²single pass flush followed by 0.5 L/ft² total recirculation for 30 min.This was followed by 1 L/ft² PW flush. This was followed by a totalrecirculation of 0.5 L/ft² 1.4% Minncare solution for 30 min. The systemwas again flushed with 1 L/ft² PW and finally recirculated for 15 minwith 0.1 N NaOH and stored.

Example 11

High Concentration Formulation of rhuMAb anti-CD20 A pilot scale UFsystem was used to concentrate and formulate rhuMAb anti-CD20 (2H7) foruse in a human phase I clinical study in a GMP manufacturing facility.Example 10 was repeated with the following exceptions.

The protein pool resulting from a preceding Q-Sepharose chromatographystep was measured to be 3.729 g 2H7/L and had a volume of 262 L. Theprotein was in a solution of 6 mM HEPES free acid/19 mM HEPES sodiumsalt and 25 mM sodium acetate that had been pH adjusted to 5.3 with 0.5M acetic acid. Just prior to the run, the temperature set point of thispool was set to 45° C. To begin the run the protein pool was transferredto the recirculation vessel through a 0.22 microns sterilizing gradefilter to a level of about 40 L in the tank. In the vessel the pool wasagitated via the top mounted impeller and the temperature was maintainedat 40-50° C.

During the UF1 the pool was concentrated to about 50 g 2H7/L (about 20L). FIG. 24 displays the value trends over time for recycle tank level(210) scaled from −0.713963 to 295.989, retentate dP (2420) scaled from−0.237899 to 98.6629, feed flow rate (250) scaled from −0.356981 to147.994, and filtrate flow rate (2450) scaled from −0.118994 to 49.3315during the process. The pool was then diafiltered with 10 diavolumes ofdiafiltration buffer. During diafiltration the temperature wasmaintained between 40 and 50° C. The diafiltration was performed at aconstant volume, which was achieved by matching the flow rate of bufferbeing transferred into the recirculation tank to the flow rate offiltrate being removed from the system. At the end of the diafiltration,the pool was further concentrated in UF2 mode to a final concentrationtarget set point of 190 g 2H7/L (5.25 L). Note in FIG. 24 theincorporation of constant dP to 40 psig control at the end of thisphase. This phase was also performed at an elevated temperature setpoint of 45° C.+/−5° C. Next, a low pressure drop mixing was performedwhere the feed pump is controlled to maintain a 20 psig pressure dropacross the feed channel. A sample was pulled and a density measurementwas performed to confirm the concentration prior to recovery. Theconcentration of this sample was 194 g 2H7/L. Table 18 summarizes thethroughput and flux results.

TABLE 18 Normalized Normalized throughput Flux Process Phase (g/ft²/hr)(LMH/psig) UF1 51 3.8 DF 46 2.2 UF2 286 1.6

Just prior to product recovery, a 30 mL sample was pulled and submittedfor detection and titer of bioburden. The results were negative (i.e.,<0.13 CFU/mL). The protein pool was recovered by the series of steps ofExample 10. The volumes of the recovered pool, diluted pool, andconditioned pool were then analyzed for protein concentration. Table 19presents the results. The membrane was regenerated as in Example 10.

TABLE 19 Concentration Yield or Volume (L) (g/L) Mass (g) {Loss} (%)Q-Pool 262 3.72 977 100 Recovered 5.0 174.0 863.0 88.3 Pool Diluted5.421 149.6 811.0 83.0 Pool

Example 12

High Concentration Formulation of rhuMAb anti-CD20 GMP Example 11 wasrepeated with the following exceptions. The protein pool resulting froma preceding Q-Sepharose chromatography step was measured to be 5.106 g2H7/L and had a volume of 196 L. The protein was in a solution of 6 mMHEPES free acid/19 mM HEPES sodium salt and 25 mM sodium acetate thathad been pH adjusted to 5.3 with 0.5 M acetic acid. Just prior to therun, the temperature setpoint of this pool was set to 45° C. To beginthe run the protein pool was transferred to the recirculation vesselthrough a 0.22 microns sterilizing grade filter to a level of about 40 Lin the tank. In the vessel the pool was agitated via the top mountedimpeller and the temperature was maintained at 40-50° C.

During the UF1 the pool was concentrated to about 50 g 2H7/L (about 20L). FIG. 25 displays the value trends over time for recycle tank level(210) scaled from 0 to 300, retentate dP (2520) scaled from 0 to 100,feed flow rate (250) scaled from 0 to 150, and filtrate flow rate (2550)scaled from 0-50 during the process. The pool was diafiltered with 10diavolumes (10×) of diafiltration buffer. During diafiltration thetemperature was maintained between 40 and 50° C. The diafiltration wasperformed at a constant volume that was achieved by matching the flowrate of buffer being transferred into the recirculation tank to the flowrate of filtrate being removed from the system. At the end of thediafiltration, the pool was further concentrated in UF2 mode to a finalconcentration target setpoint of 190 g 2H7/L (5.26 L) again utilizingconstant dP control at the very end of this phase (see FIG. 25). Thisphase was also performed at an elevated temperature set point of 45°C.+/−5° C. Next, a low pressure drop mixing was performed where the feedpump is controlled to maintain a 20 psig pressure drop across the feedchannel. A sample was pulled and a density measurement was performed toconfirm the concentration prior to recovery. The concentration of thissample was 191 g 2H7/L. Table 20 summarizes the throughput and fluxresults.

TABLE 20 Normalized Normalized throughput Flux Process Phase (g/ft²/hr)(LMH/psig) UF1 67 3.6 DF 47 2.1 UF2 292 1.8

Just prior to product recovery, a 30 mL sample was pulled and submittedfor detection and titer of bioburden. The results were negative (i.e.,<0.13 CFU/mL). The protein pool was recovered by a series of steps as inExample 11. The volumes of the recovered pool, diluted pool, andconditioned pool were then analyzed for protein concentration. Table 21presents the results. The membrane was regenerated as in Example 1 l.

TABLE 21 Concentration Yield or Volume (L) (g/L) Mass (g) {Loss} (%)Q-Pool 196 5.106 1000 100 Recovered 4.9 187.1 918.0 91.8 Pool DilutedPool 6.075 150.9 916.9 91.7

All publications, patents, and patent documents are incorporated byreference herein in their entirety, as though individually incorporatedby reference. The disclosure has been described with reference tovarious specific and preferred embodiments and techniques. However, itshould be understood that many variations and modifications can be madewhile remaining within the spirit and scope of the disclosure.

1. A process for preparing highly concentrated antibody compositionscomprising: a first ultrafiltering of a first antibody preparation toprovide a second antibody preparation; a diafiltering the secondantibody preparation to provide a diafiltered intermediate antibodypreparation; and a second ultrafiltering of the diafiltered intermediateantibody preparation to provide a third antibody preparation; whereinone or more of the first ultrafiltering, second ultrafiltering, and thediafiltering are accomplished at about 30° C. to about 50° C.
 2. Theprocess of claim 1 wherein one or more of the first ultrafiltering, thesecond ultrafiltering, and the diafiltering are accomplished at fromabout 35° C. to about 50° C.
 3. The process of claim 1 wherein one ormore of the first ultrafiltering, the second ultrafiltering, and thediafiltering are accomplished at from about 45° C.
 4. The process ofclaim 1 wherein one or more of the first ultrafiltering, the secondultrafiltering, and the diafiltering are accomplished at 45° C. plus orminus 5° C.
 5. The process of claim 1 wherein the first antibodypreparation has an antibody concentration of from about 0.1 to about 10grams per liter.
 6. The process of claim 1 wherein the first antibodypreparation has an antibody concentration of about 1 to about 5 gramsper liter.
 7. The process of claim 1 wherein the second antibodypreparation has an antibody concentration of from about 10 to about 50grams per liter.
 8. The process of claim 1 wherein the second antibodypreparation has an antibody concentration of about 20 to about 40 gramsper liter.
 9. The process of claim 1 wherein the third antibodypreparation has an antibody concentration of from about 50 to about 250grams per liter.
 10. The process of claim 1 wherein the third antibodypreparation has an antibody concentration of from about 100 to about 230grams per liter.
 11. The process of claim 1 wherein the third antibodypreparation has an antibody concentration of from about 170 to about 200grams per liter.
 12. The process of claim 1 wherein the diafilteredintermediate antibody preparation and the third antibody preparationcomprise the ultra-filter retentate.
 13. The process of claim 1 whereinthe intermediate antibody preparation has an antibody concentration ofabout 25 to about 35 grams per liter and the third antibody preparationhas an antibody concentration of from about 170 to about 200 grams perliter.
 14. The process of claim 1 wherein the antibody preparationcomprises anti-IgE antibodies.
 15. The process of claim 1 wherein theprocess is accomplished in from about 1 to 10 hours.
 16. The process ofclaim 1 wherein the process is accomplished in from about 2 to 5 hours.17. The process of claim 1 wherein the process is accomplished in about3 hours.
 18. The process of claim 1 wherein the first and the secondultrafiltering are accomplished with an ultra-filter membrane having anominal pore size of about 5 to about 50 kilo Daltons.
 19. The processof claim 1 wherein the first and the second ultrafiltering areaccomplished with an ultra-filter membrane having a nominal pore size ofabout 10 to about 30 kilo Daltons.
 20. The process of claim 1 whereinthe first antibody preparation contains an antibody having an apparentmolecular weight of about 100 to about 200 kilo Daltons.
 21. The processof claim 1 wherein the first antibody preparation contains an antibodyhaving an apparent molecular weight of about 150 kilo Daltons.
 22. Theprocess of claim 1 wherein the first ultrafiltering concentrates thefirst antibody preparation to provide the second antibody preparationhaving an antibody concentration of about 30 grams per liter and thesecond ultrafiltering concentrates the intermediate antibody preparationto provide the third antibody preparation having an antibodyconcentration of about 170 to about 200 grams per liter.
 23. The processof claim 1 wherein the first ultrafiltering and the secondultrafiltering are accomplished with the same ultra-filter membrane. 24.The process of claim 1 wherein the first ultrafiltering and the secondultrafiltering are accomplished with a regenerated cellulose compositeultra-filter membrane.
 25. The process of claim 1 wherein thediafiltering accomplishes a buffer exchange at constant volume, constantconcentration, or both.
 26. The process of claim 1 wherein thediafiltering accomplishes a buffer exchange of from about 5 to about 15fold volumes.
 27. The process of claim 1 wherein the diafilteringaccomplishes a buffer exchange of from about 8 fold volumes.
 28. Theprocess of claim 1 wherein the diafiltering exchanges a first buffer fora second buffer.
 29. The process of claim 28 wherein the first buffercomprises a mixture of aqueous sodium chloride and a TRIS buffer, andthe second buffer comprises a mixture of aqueous histidine chloride andarginine chloride.
 30. The process of claim 1 wherein the firstultrafiltering, the second ultrafiltering, and the diafiltering areaccomplished with tangential flow filtration across an ultra-filtermembrane.
 31. The process of claim 1 wherein the first ultrafiltering,the second ultrafiltering, and the diafiltering are accomplished withtangential flow filtration across the same ultra-filter membrane. 32.The process of claim 1 wherein the yield of the third antibodypreparation is greater than about 70 weight % based on the weight ofantibodies in the first antibody preparation.
 33. The process of claim32 wherein the yield of the third antibody preparation is from about 80to about 100 weight % based on the weight of antibodies in the firstantibody preparation.
 34. The process of claim 1 wherein the firstultrafiltering has a recirculation rate of from about 0.5 L/min/ft² toabout 5 L/min/ft².
 35. The process of claim 1 wherein the ultrafilteringand diafiltering are accomplished at a transmembrane pressure of fromabout 10 to about 50 p.s.i.
 36. The process of claim 1 wherein there isprovided an antibody concentrate with a detectable bio-burden of lessthan about 100 CFU/mL.
 37. A process for concentrating proteinscomprising: a first ultrafiltering of a first protein mixture to providea second protein mixture; a diafiltering of the second protein mixtureto provide a diafiltered protein mixture; and a second ultrafiltering ofthe diafiltered protein mixture to provide a third protein mixture,wherein one or more of the first ultrafiltering, the diafiltering, andthe second ultrafiltering are accomplished at from about 30° C. to about50° C.